Hydrocracking Catalyst Containing Beta and Y Zeolites, and Process for its use to make Distillate

ABSTRACT

Increased selectivity of middle distillate and/or increased catalyst activity are obtained in a hydrocracking process by the use of a catalyst containing a beta zeolite and a Y zeolite having a unit cell size of 24.33 to 24.38 angstrom. The catalyst may also contain an additional Y zeolite having a unit cell size of from 24.25 to 24.32 angstrom.

FIELD OF THE INVENTION

The invention relates to catalyst compositions and their use inhydrocarbon conversion processes, particularly hydrocracking. Theinvention more specifically relates to a catalyst composition thatcomprises a Y zeolite and a beta zeolite as active cracking components.The invention specifically relates to a hydrocracking process thatproduces middle distillate.

BACKGROUND OF THE INVENTION

Petroleum refiners often produce desirable products such as turbinefuel, diesel fuel, and other hydrocarbon liquids known as middledistillates, as well as lower boiling liquids such as naphtha andgasoline, by hydrocracking a hydrocarbon feedstock derived from crudeoil. Hydrocracking also has other beneficial results such as removingsulfur and nitrogen from the feedstock by hydrotreating. Feedstocks mostoften subject to hydrocracking are gas oils and heavy gas oils recoveredfrom crude oil by distillation.

Hydrocracking is generally carried out by contacting, in an appropriatereactor vessel, the gas oil or other hydrocarbon feedstock with asuitable hydrocracking catalyst under appropriate conditions, includingan elevated temperature and an elevated pressure and the presence ofhydrogen so as to yield a lower overall average boiling point productcontaining a distribution of hydrocarbon products desired by therefiner. Although the operating conditions within a hydrocrackingreactor have some influence on the yield of the products, thehydrocracking catalyst is a prime factor in determining such yields.

Hydrocracking catalysts are subject to initial classification on thebasis of the nature of the predominant cracking component of thecatalyst. This classification divides hydrocracking catalysts into thosebased upon an amorphous cracking component such as silica-alumina andthose based upon a zeolitic cracking component such as beta or Yzeolite. Hydrocracking catalysts are also subject to classification onthe basis of their intended predominant product of which the two mainproducts are naphtha and “distillate”, a term which in the hydrocrackingrefining art refers to distillable petroleum derived fractions having aboiling point range that is above that of naphtha. Distillate typicallyincludes the products recovered at a refinery as kerosene and dieselfuel. At the present time, distillate is in high demand. For thisreason, refiners have been focusing on hydrocracking catalysts whichselectively produce a distillate fraction.

The three main catalytic properties by which the performance of ahydrocracking catalyst for making distillate is evaluated are activity,selectivity, and stability. Activity may be determined by comparing thetemperature at which various catalysts must be utilized under otherwiseconstant hydrocracking conditions with the same feedstock so as toproduce a given percentage, normally about 65 percent, of productsboiling in the desired range, e.g., below 371° C. (700° F.) fordistillate. The lower the temperature required for a given catalyst, themore active such a catalyst is in relation to a catalyst requiring ahigher temperature. Selectivity of hydrocracking catalysts may bedetermined during the foregoing described activity test and is measuredas a percentage of the fraction of the product boiling in the desireddistillate product range, e.g., from 149° C. (300° F.) to 371° C. (700°F.). Stability is a measure of how well a catalyst maintains itsactivity over an extended time period when treating a given hydrocarbonfeedstock under the conditions of the activity test. Stability isgenerally measured in terms of the change in temperature required perday to maintain a 65 percent or other given conversion.

Although cracking catalysts for producing distillate are known and usedin commercial environments, there is always a demand for newhydrocracking catalysts with superior selectivity at a given activityand/or superior activity at a given selectivity for producingdistillate.

BRIEF SUMMARY OF THE INVENTION

It has been found that hydrocracking catalysts containing a Y zeolitehaving a unit cell size or dimension a_(o) of from 24.33 to 24.38angstrom (hereinafter Y Zeolite II) and containing a beta zeolitepreferably having an overall silica to alumina (SiO₂ to Al₂O₃) moleratio of less than 30 and a SF₆ adsorption capacity of at least 28weight-percent (hereinafter wt-%) have substantially improvedselectivity at a given activity or substantially improved activity at agiven selectivity compared to other hydrocracking catalysts nowcommercially available for use in hydrocracking processes for producingdistillate. The catalyst also contains a metal hydrogenation componentsuch as nickel, cobalt, tungsten, molybdenum, or any combinationthereof. The catalyst contains from 0.5 to 5 wt-% beta zeolite based onthe combined weight of the beta zeolite, the Y Zeolite II, and thesupport on a dried basis, and the catalyst has a weight ratio of the YZeolite II to the beta zeolite of from 0.5 to 5 on a dried basis. The YZeolite II has an overall silica to alumina mole ratio of from 5.0 to11.0. In one embodiment, the catalyst contains an additional Y zeoliteand the Y Zeolite II has a higher unit cell size than that of theadditional zeolite.

It is believed that a hydrocracking catalyst containing such a Y zeoliteand such a beta zeolite is novel to the art.

Under typical hydrocracking conditions, including elevated temperatureand pressure and the presence of hydrogen, such catalysts are highlyeffective for converting gas oil and other hydrocarbon feedstocks to aproduct of lower average boiling point and lower average molecularweight. In one embodiment, the product contains a relatively largeproportion of components boiling in the distillate range, which asdefined herein is from 149° C. (300° F.) to 371° C. (700° F.).

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a graph of distillate selectivity versus relative catalystactivity for several hydrocracking catalysts. FIG. 2 is a graph of theratio of heavy distillate selectivity to light distillate selectivityversus relative catalyst activity for several hydrocracking catalysts.

INFORMATION DISCLOSURE

Beta and Y zeolites have been proposed in combination as components ofseveral different catalysts including catalysts for hydrocracking. Forinstance, U.S. Pat. No. 5,275,720; U.S. Pat. No. 5,279,726; and U.S.Pat. No. 5,350,501 describe hydrocracking processes using a catalystcomprising a beta zeolite and a Y zeolite. U.S. Pat. No. 5,350,501describes a hydrocracking process using a catalyst comprising, amongother components, zeolite beta and a Y zeolite having a unit cell sizebetween 24.25 and 24.35 angstrom and a water vapor sorption capacity at4.6 mm water vapor partial pressure and 25° C. less than 8.0 percent byweight of the zeolite. US-A1-2004/0152587 describes a hydrocrackingcatalyst comprising a carrier comprising a zeolite of the faujasitestructure having a unit cell size in the range of from 24.10 to 24.40angstrom, a bulk silica to alumina ratio above about 12, and a surfacearea of at least about 850 m²/g, and the catalyst may contain a secondzeolite such as beta zeolite, ZSM-5 zeolite, or a Y zeolite of adifferent unit cell size. Two different Y zeolites have also beenproposed in combination as components of several different catalystsincluding catalysts for hydrocracking, as described in U.S. Pat. No.4,661,239 and U.S. Pat. No. 4,925,546.

DETAILED DESCRIPTION OF THE INVENTION

The process and composition disclosed herein may be used to convert afeedstock containing organic compounds into products, particularly byacid catalysis, such as hydrocracking organic compounds especiallyhydrocarbons into a product of lower average boiling point and loweraverage molecular weight. The composition, which may be a catalystand/or a catalyst support, comprises a beta zeolite and Y Zeolite II.The composition may also comprise a refractory inorganic oxide. Whenused as a catalyst for hydrocracking, the composition contains a betazeolite, Y Zeolite II, a refractory inorganic oxide, and a hydrogenationcomponent.

The hydrocracking process and composition disclosed herein centers onusing a catalyst containing a particular beta zeolite and a particular Yzeolite. The composition may optionally contain an additional Y zeolite.The beta zeolite preferably has a relatively low silica to alumina moleratio and a relatively high SF₆ adsorption capacity. Y Zeolite II has aunit cell size of from 24.33 to 24.38 angstrom. When the additionalzeolite is present, Y Zeolite II has a higher unit cell size than thatof Y Zeolite I. It has been found that differing performance resultswhen such a beta zeolite and such Y zeolites are incorporated in ahydrocracking catalysts in this way. Compared to catalysts containingone or two Y zeolites, the selectivity of product boiling in thedistillate range is higher at a given activity or the activity is higherat a given selectivity of product boiling in the distillate range.

Beta zeolite is well known in the art as a component of hydrocrackingcatalysts. Beta zeolite is described in U.S. Pat. No. 3,308,069 and U.S.Reissue No. 28341, which are hereby incorporated by reference herein intheir entireties. The beta zeolite that is used in the process andcomposition disclosed herein has a silica to alumina mole ratio of lessthan 30 in one embodiment, less than 25 in another embodiment, more than9 and less than 30 in yet another embodiment, more than 9 and less than25 in a further embodiment, more than 20 and less than 30 in anotherembodiment, or more than 15 and less than 25 in still anotherembodiment. As used herein, unless otherwise indicated, the silica toalumina (SiO₂ to Al₂O₃) mole ratio of a zeolite is the mole ratio asdetermined on the basis of the total or overall amount of aluminum andsilicon (framework and non-framework) present in the zeolite, and issometimes referred to herein as the overall silica to alumina (SiO₂ toAl₂O₃) mole ratio.

Beta zeolite is usually synthesized from a reaction mixture containing atemplating agent. The use of templating agents for synthesizing betazeolite is well known in the art. For example, U.S. Pat. No. 3,308,069and U.S. Reissue No. 28341 describe using tetraethylammonium hydroxideand U.S. Pat. No. 5,139,759, which is hereby incorporated herein byreference in its entirety, describes using the tetraethylammonium ionderived from the corresponding tetraethylammonium halide. Anotherstandard method of preparing beta zeolite is described in the booktitled Verified Synthesis of Zeolitic Materials, by H. Robson (editor)and K. P. Lillerud (XRD Patterns), second revised edition, ISBN0-444-50703-5, Elsevier, 2001. It is believed that the choice of aparticular templating agent is not critical to the success of theprocess disclosed herein. In one embodiment the beta zeolite is calcinedin air at a temperature of from 500 to 700° C. (932 to 1292° F.) for atime sufficient to remove to remove the templating agent from the betazeolite. Calcination to remove the templating agent can be done beforeor after the beta zeolite is combined with the support and/or thehydrogenation component. Although it is believed that the templatingagent could be removed at calcination temperatures above 700° C. (1292°F.), very high calcination temperatures could significantly decrease theSF₆ adsorption capacity of beta zeolite. For this reason it is believedthat calcination temperatures above 750° C. (1382° F.) for removing thetemplating agent should be avoided when preparing the beta zeolite foruse in the process disclosed herein. It is critical to the processdisclosed herein that the SF₆ adsorption capacity of the beta zeolite isat least 28 wt-%.

While it is known that steaming a zeolite such as beta results inchanges to the actual crystalline structure of the zeolite, theabilities of present day analytical technology have not made it possibleto accurately monitor and/or characterize these changes in terms ofimportant structural details of the zeolite. Instead, measurements ofvarious physical properties of the zeolite such as surface area are usedas indicators of changes that have occurred and the extent of thechanges. For instance, it is believed that a reduction in the zeolite'scapacity to adsorb sulfur hexafluoride (SF₆) after being steamed iscaused by a reduction in the crystallinity of the zeolite or in the sizeor accessibility of the zeolite's micropores. It is, however, anindirect correlation of the changes in the zeolite that may beundesirable, since the SF₆ adsorption capacity in the catalyst used inthe process and composition disclosed herein is relatively high. Inembodiments of the process and composition disclosed herein, the SF₆adsorption capacity of the beta zeolite, whether steam treated or not,should be at least 28 wt-%.

Accordingly, the beta zeolite of the process and composition disclosedherein may be characterized in terms of SF₆ adsorption. This is arecognized technique for the characterization of microporous materialssuch as zeolites. It is similar to other adsorption capacitymeasurements, such as water capacity, in that it uses weight differencesto measure the amount of SF₆ which is adsorbed by a sample which hasbeen pretreated to be substantially free of the adsorbate. SF₆ is usedin this test since because its size and shape hinders its entrance intopores having a diameter of less than about 6 angstrom. It thus can beused as one measurement of available pore mouth and pore diametershrinkage. This in turn is a measurement of the effect of steaming onthe zeolite. In a simplistic description of this measurement method, thesample is preferably first predried in a vacuum at 300° C. (572° F.) forone hour, then heated at atmospheric pressure in air at 650° C. (1202°F.) for two hours, and finally weighed. It is then exposed to the SF₆for one hour while the sample is maintained at a temperature of 20° C.(68° F.). The vapor pressure of the SF₆ is maintained at that providedby liquid SF₆ at 400 torr (53.3 kPa (7.7 psi)). The sample is againweighed to measure the amount of adsorbed SF₆. The sample may besuspended on a scale during these steps to facilitate these steps.

In any mass production procedure involving techniques such as steamingand heating there is a possibility for individual particles to besubjected to differing levels of treatment. For instance, particles onthe bottom of a pile moving along a rotary kiln may not be subjected tothe same atmosphere or temperature as the particles which cover the topof the pile. This factor must be considered during manufacturing andalso during analysis and testing of the finished product. It is,therefore, recommended that any test measure done on the material isperformed on a representative composite sample of the entire quantity offinished product to avoid being misled by measurements performed onindividual particles or on a non-representative sample. For instance, anadsorption capacity measurement is made on a representative compositesample.

Although the process and the composition disclosed herein can use a betazeolite that has not been subjected to a steaming treatment, the processand the composition disclosed herein can also use beta zeolite that issubjected to steaming, provided that the steaming is relatively mild incomparison to steaming of beta zeolite in the literature. Under theproper conditions and for the proper time, steaming beta zeolite canyield a catalyst that can be used in the process and compositiondisclosed herein.

Hydrothermally treating zeolites for use in hydrocracking catalysts is arelatively blunt tool. For any given zeolite, steaming decreases theacidity of the zeolite. When the steamed zeolite is used as ahydrocracking catalyst, the apparent result is that the overalldistillate yield increases but the catalyst's activity decreases. Thisapparent tradeoff between yield and activity has meant that to achievehigh activity means not to steam the beta zeolite, but at the expense oflower product yields. This apparent tradeoff between yield and activitymust be considered and is a limit to the improvement that appears to beobtainable by steaming the beta zeolite. When the steamed beta zeoliteis used in the catalysts disclosed herein, the improvement in activityover catalysts containing only Y zeolite would appear limited while theimprovement in yield over such catalysts would appear more enhanced.

If the beta zeolite is to be steamed, such steaming can be performedsuccessfully in different ways, with the method which is actuallyemployed commercially often being greatly influenced and perhapsdictated by the type and capability of the available equipment. Steamingcan be performed with the beta zeolite retained as a fixed mass or withthe beta zeolite being confined in a vessel or being tumbled whileconfined in a rotating kiln. The important factors are uniform treatmentof all beta zeolite particles under appropriate conditions of time,temperature and steam concentration. For instance, the beta zeoliteshould not be placed such that there is a significant difference in theamount of steam contacting the surface and the interior of the betazeolite mass. The beta zeolite may be steam treated in an atmospherehaving live steam passing through the equipment providing low steamconcentration. This may be described as being at a steam concentrationof a positive amount less than 50 mol-%. Steam concentrations may rangefrom 1 to 20 mol-% or from 5 to 10 mol-%, with small-scale laboratoryoperations extending toward higher concentrations. The steaming may beperformed for a positive time period of less than or equal to 1 or 2hours or for 1 to 2 hours at a temperature of less than or equal toabout 600° C. (1112° F.) at atmospheric pressure and a positive contentof steam of less than or equal to 5 mol-%. The steaming may be performedfor a positive time period of less than or equal to 2 hours at atemperature of less than or equal to about 650° C. (1202° F.) atatmospheric pressure and a positive content of steam of less than orequal to 10 mol-%. The steam contents are based on the weight of vaporscontacting the beta zeolite. Steaming at temperatures above 650° C.(1202° F.) appears to result in beta zeolite that is not useful in theprocess disclosed herein since the SF₆ adsorption capacity of theresulting beta zeolite is too low. Temperatures below 650° C. (1202° F.)can be used, and the steaming temperature can be from about 600° C.(1112° F.) to about 650° C. (1202° F.), or less than 600° C. (1112° F.).It is taught in the art that there is normally an interplay between timeand temperature of steaming, with an increase in temperature reducingthe required time. Nevertheless, if steaming is done, for good resultsit appears a time period of about ½ to about 2 hours or about 1 to about1½ hours can be used. The method of performing steaming on a commercialscale may be by means of a rotary kiln having steam injected at a ratewhich maintains an atmosphere of about 10 mol-% steam.

An exemplary lab scale steaming procedure is performed with the zeoliteheld in a 6.4 cm (2½ inch) quartz tube in a clam shell furnace. Thetemperature of the furnace is slowly ramped up by a controller. Afterthe temperature of the zeolite reaches 150° C. (302° F.) steam generatedfrom deionized water held in a flask is allowed to enter the bottom ofthe quartz tube and pass upward. Other gas can be passed into the tubeto achieve the desired steam content. The flask is refilled as needed.In the exemplary procedure the time between cutting in the steam and thezeolite reaching 600° C. (1112° F.) is about one hour. At the end of theset steam period the temperature in the furnace is reduced by resettingthe controller to 20° C. (68° F.). The furnace is allowed to cool to400° C. (752° F.) (about 2 hours) and the flow of steam into the quartztube is stopped. The sample is removed at 100° C. (212° F.) and placedin a lab oven held overnight at 110° C. (230° F.) with an air purge.

The beta zeolite of the process and composition disclosed herein is nottreated with an acid solution to effect dealumination. In this regard itis noted that essentially all raw (as synthesized) beta zeolite isexposed to an acid to reduce the concentration of alkali metal (e.g.,sodium) which remains from synthesis. This step in the beta zeolitemanufacture procedure is not considered part of the treatment ofmanufactured beta zeolite as described herein. In one embodiment, duringthe treatment and catalyst manufacturing procedures, the beta zeolite isexposed to an acid only during incidental manufacturing activities suchas peptization during forming or during metals impregnation. In anotherembodiment, the beta zeolite is not acid washed after the steamingprocedure as to remove aluminum “debris” from the pores.

Also included in the process and composition disclosed herein is a Yzeolite having a unit cell size of from 24.33 to 24.38 angstrom. This Yzeolite is sometimes referred to herein as Y Zeolite II in order todistinguish this Y zeolite from an optional additional Y zeolite havinga different unit cell size and described hereinafter. Y Zeolite IIpreferably has a unit cell size of from 24.34 to 24.36 angstrom. YZeolite II can have an overall silica to alumina mole ratio of from 5.0to 12.0 in one embodiment, from 5.0 to 1.0 in another embodiment, andfrom 5.0 to 10.0 in yet another embodiment. The process and compositiondisclosed herein require a Y Zeolite II.

Optionally and in addition to Y Zeolite II, the process and compositiondisclosed may include an additional Y zeolite, which is sometimesreferred to herein as Y Zeolite I. Y Zeolite I has a different unit cellsize from the unit cell size of Y Zeolite II. The unit cell size of theY Zeolite I is preferably at least 0.04 angstrom smaller than the unitcell size of Y Zeolite II. The unit cell size of Y Zeolite I is morepreferably from 24.25 to 24.32 angstrom, and even more preferably from24.26 to 24.30 angstrom. Y Zeolite I can have an overall silica toalumina mole ratio of from 5.0 to 12.0 in one embodiment, from 5.0 to11.0 in another embodiment, and from 5.0 to 10.0 in yet anotherembodiment.

The option of adding Y Zeolite I during the manufacturing process givescatalyst producers flexibility to make products that meet the individualrequirements of hydrocracking unit operators. The presence of Y ZeoliteI in the catalyst changes the properties of the catalyst without theneed to change how the Y Zeolite II itself is prepared or the amount ofY Zeolite II used in the catalyst. In some instances, however, adding YZeolite I decreases the requirement for Y Zeolite II, which is anadditional advantage when sufficient quantities of Y Zeolite II are notavailable. Hydrocracking unit operators, especially those producingdistillate, can use catalysts containing both Y Zeolite I and Y ZeoliteII as a tool to satisfy their particular and sometimes uniquerequirements for hydrocracking catalyst activity and selectivity.

The term “Y zeolite” as used herein is meant to encompass allcrystalline zeolites having either the essential X-ray powderdiffraction pattern set forth in U.S. Pat. No. 3,130,007 or a modified Yzeolite having an X-ray powder diffraction pattern similar to that ofU.S. Pat. No. 3,130,007 but with the d-spacings shifted somewhat due, asthose skilled in the art will realize, to cation exchanges,calcinations, etc., which are generally necessary to convert the Yzeolite into a catalytically active and stable form. Y Zeolite I and YZeolite II are modified Y zeolites in comparison to the Y zeolite taughtin U.S. Pat. No. 3,130,007. As used herein, unit cell size means theunit cell size as determined by X-ray powder diffraction.

The Y zeolites used in the process and composition disclosed herein arelarge pore zeolites having an effective pore size greater than 7.0angstrom. Since some of the pores of the Y zeolites are relativelylarge, the Y zeolites allow molecules relatively free access to theirinternal structure. The pores of the Y zeolites permit the passagethereinto of benzene molecules and larger molecules and the passagetherefrom of reaction products.

One group of Y zeolites that may be used in the process and compositiondisclosed herein as Y Zeolite I, Y Zeolite II, or both includes zeolitesthat are sometimes referred to as ultrastable or ultrahydrophobic Yzeolites. The composition and properties of this group of Y zeolitesare, in essence, prepared by a four step procedure. First, a Y zeolitein the alkali metal form (usually sodium) and typically having a unitcell size of about 24.65 angstrom is cation exchanged with ammoniumions. The ammonium exchange step typically reduces the sodium content ofthe starting sodium Y zeolite from a value usually greater than about 8wt-%, usually from about 10 to about 13 wt-%, calculated as Na₂O, to avalue in the range from about 0.6 to about 5 wt-%, calculated as Na₂O.Methods of carrying out the ion exchange are well known in the art.

Second, the Y zeolite from the first step is calcined in the presence ofwater vapor. For example, the Y zeolite is calcined in the presence ofat least 1.4 kPa(absolute) (hereinafter kPa(a)) (0.2 psi(absolute)(hereinafter psi(a))), at least 6.9 kPa(a) (1.0 psi(a)), or at least 69kPa(a) (10 psi(a)) water vapor, in three embodiments. In two otherembodiments, the Y zeolite is calcined in an atmosphere consistingessentially of or consisting of steam. The Y zeolite is calcined so asto produce a unit cell size in the range of 24.40 to 24.64 angstrom.

Third, the Y zeolite from the second step is ammonium exchanged onceagain. The second ammonium exchange further reduces the sodium contentto less than about 0.5 wt-%, usually less than about 0.3 wt-%,calculated as Na₂O.

Fourth, the Y zeolite from the third step is treated further so as toyield Y zeolite having a unit cell size from 24.25 to 24.32 angstrom orpreferably from 24.26 to 24.30 angstrom in the case of Y Zeolite I. Inthe case of Y Zeolite II, the treatment yields a Y zeolite having a unitcell size from 24.33 to 24.38 angstrom or preferably from 24.34 to 24.36angstrom. The zeolite Y resulting from the fourth step has an overallsilica to alumina mole ratio from 5.0 to 12.0 in one embodiment, from5.0 to 11.0 in another embodiment, and from 5.0 to 10.0 in yet anotherembodiment. The treatment of the fourth step can comprise any of thewell known techniques for dealuminating zeolites in general andultrastable Y zeolite in particular so as to yield the desired unit cellsize and overall silica to alumina mole ratio. The fourth treatment stepmay change the unit cell size and/or the framework silica to aluminamole ratio, with or without changing the overall silica to alumina moleratio. Generally, zeolite dealumination is accomplished by chemicalmethods such as treatments with acids, e.g., HCl, with volatile halides,e.g., SiCl₄, or with chelating agents such as ethylenediaminetetraaceticacid (EDTA). Another common technique is a hydrothermal treatment of thezeolite in either pure steam or in air/steam mixtures, preferably suchas calcining in the presence of sufficient water vapor (for example, inan atmosphere consisting essentially of steam, and most preferablyconsisting of steam) so as to yield the desired unit cell size andoverall silica to alumina mole ratio.

The above-discussed preparation procedure for Y zeolites used in theprocess and composition disclosed herein differs from the procedure forthe Y zeolites taught in U.S. Pat. No. 3,929,672 by the addition of thefourth treatment step. U.S. Pat. No. 3,929,672, which is herebyincorporated herein by reference in its entirety, discloses a method fordealuminating an ultrastable Y zeolite. U.S. Pat. No. 3,929,672 teachesa preparation procedure wherein a sodium Y zeolite is partiallyexchanged with ammonium ions, followed by steam calcination undercontrolled temperature and steam partial pressure, followed by yetanother ammonia exchange and then by an optional calcination step in adry atmosphere. The exchange and steam calcination steps can be repeatedto achieve the desired degree of dealumination and unit cell sizereduction. The zeolites of U.S. Pat. No. 3,929,672 are known under thedesignation Y-84 or LZY-84 commercially available from UOP LLC, DesPlaines, Ill., U.S.A. Y-84 or LZY-84 zeolites may be produced by thefirst three steps just mentioned, but optionally one may include afurther calcination step in a dry atmosphere, e.g., a calcination inwater- and steam-free air, at 482° C. (900° F.) or higher.

The above-discussed preparation procedure for Y zeolites used in theprocess and composition disclosed herein is similar to the procedure forthe Y zeolites taught in U.S. Pat. No. 5,350,501. However, particularconditions in the above-discussed fourth treatment step are selected inorder to produce critical ranges of unit cell size for Y Zeolite II andoptional Y Zeolite I. U.S. Pat. No. 5,350,501, which is herebyincorporated herein by reference in its entirety, discloses a fourthstep that involves calcining the resulting zeolite from the thirdtreatment step in the presence of sufficient water vapor (in anatmosphere consisting essentially of steam or consisting of steam) so asto yield a unit cell size below 24.40, and most preferably no more than24.35 angstrom, and with a relatively low sorptive capacity for watervapor. The Y zeolite produced by the four-step procedure in U.S. Pat.No. 5,350,501 is a UHP-Y zeolite, an ultrahydrophobic Y zeolite asdefined in U.S. Pat. No. 5,350,501. U.S. Pat. No. 5,350,501 defines a“UHP-Y” zeolites as zeolite aluminosilicates having among otherproperties, a unit cell size or dimension as of less than 24.45 angstromand a sorptive capacity for water vapor at 25° C. and a p/p_(o) value of0.10 of less than 10.00 weight percent. The most preferred UHP-Y zeolitein U.S. Pat. No. 5,350,501 is LZ-10.

Another group of Y zeolites which may be used in the process andcomposition disclosed herein as Y Zeolite I, Y Zeolite II, or both maybe prepared by dealuminating a Y zeolite having an overall silica toalumina mole ratio below about 5 and are described in detail in U.S.Pat. No. 4,503,023; U.S. Pat. No. 4,597,956 and U.S. Pat. No. 4,735,928,which are hereby incorporated herein by reference in their entireties.U.S. Pat. No. 4,503,023 discloses another procedure for dealuminating aY zeolite involving contacting the Y zeolite with an aqueous solution ofa fluorosilicate salt using controlled proportions, temperatures, and pHconditions which avoid aluminum extraction without silicon substitution.U.S. Pat. No. 4,503,023 sets out that the fluorosilicate salt is used asthe aluminum extractant and also as the source of extraneous siliconwhich is inserted into the Y zeolite structure in place of the extractedaluminum. The salts have the general formula:

(A)_(2/b)SiF₆

wherein A is a metallic or nonmetallic cation other than H⁺ having thevalence “b.”Cations represented by “A” are alkylammonium, NH₄ ⁺, Mg⁺⁺,Li⁺, Na⁺, K⁺, Ba⁺⁺, Cd⁺⁺, Cu⁺⁺, H⁺, Ca⁺⁺, Cs⁺, Fe⁺⁺, Co⁺⁺, Pb⁺⁺, Mn⁺⁺,Rb⁺, Ag⁺, Sr⁺⁺, Ti⁺, and Zn⁺⁺.

A preferred member of this group is known as LZ-210, a zeoliticaluminosilicate molecular sieve commercially available from UOP LLC, DesPlaines, Ill., U.S.A. LZ-210 zeolites and the other zeolites of thisgroup are conveniently prepared from a Y zeolite starting material. TheLZ-210 zeolite has an overall silica to alumina mole ratio from 5.0 to12.0 in one embodiment, from 5.0 to 11.0 in another embodiment, and from5.0 to 10.0 in yet another embodiment. The unit cell size can bepreferably from 24.25 to 24.32 angstrom or more preferably from 24.26 to24.30 angstrom in the case of Y Zeolite I. In the case of Y Zeolite 11,the unit cell size can be from 24.33 to 24.38 angstrom or preferablyfrom 24.34 to 24.36 angstrom. The LZ-210 class of zeolites used in theprocess and composition disclosed herein have a composition expressed interms of mole ratios of oxides as in the following formula:

(0.85−1.1)M_(2/n)O:Al₂O₃ :xSiO₂

wherein “M” is a cation having the valence “n” and “x” has a value from5.0 to 12.0.

In general, LZ-210 zeolites may be prepared by dealuminating Y-typezeolites using an aqueous solution of a fluorosilicate salt, preferablya solution of ammonium hexafluorosilicate. The dealumination can beaccomplished by placing a Y zeolite, normally but not necessarily anammonium exchanged Y zeolite, into an aqueous reaction medium such as anaqueous solution of ammonium acetate, and slowly adding an aqueoussolution of ammonium fluorosilicate. After the reaction is allowed toproceed, a zeolite having an increased overall silica to alumina moleratio is produced. The magnitude of the increase is dependent at leastin part on the amount of fluorosilicate solution contacted with thezeolite and on the reaction time allowed. Normally, a reaction time ofbetween about 10 and about 24 hours is sufficient for equilibrium to beachieved. The resulting solid product, which can be separated from theaqueous reaction medium by conventional filtration techniques, is a formof LZ-210 zeolite. In some cases this product may be subjected to asteam calcination by methods well known in the art. For instance, theproduct may be contacted with water vapor at a partial pressure of atleast 1.4 kPa(a) (0.2 psi(a)) for a period of between about ¼ to about 3hours at a temperature between 482° C. (900° F.) and about 816° C.(1500° F.) in order to provide greater crystalline stability. In somecases the product of the steam calcination may be subjected to anammonium-exchange by methods well known in the art. For instance, theproduct may be slurried with water after which an ammonium salt is addedto the slurry. The resulting mixture is typically heated for a period ofhours, filtered, and washed with water. Methods of steaming andammonium-exchanging LZ-210 zeolite are described in U.S. Pat. No.4,503,023; U.S. Pat. No. 4,735,928 and U.S. Pat. No. 5,275,720.

Optional Y Zeolite I prepared by the above-discussed preparationprocedures and used in the process and composition disclosed herein havethe essential X-ray powder diffraction pattern of zeolite Y and a unitcell size or dimension a_(o) of preferably from 24.25 to 24.32 angstrom,more preferably from 24.26 to 24.30 angstrom. The Y Zeolite II preparedby the above-discussed preparation procedures and used in the processand composition disclosed herein have the essential X-ray powderdiffraction pattern of zeolite Y and a unit cell size or dimension a_(o)of from 24.33 to 24.38 angstrom, preferably from 24.34 to 24.36angstrom. Y Zeolite I, Y Zeolite II, or both can have an overall silicato alumina mole ratio of from 5.0 to 12.0 in one embodiment, from 5.0 to1.0 in another embodiment, and from 5.0 to 10.0 in yet anotherembodiment. Y Zeolite I and/or Y Zeolite II may have a surface area(BET) of at least about 500 m²/g, less than about 800 m²/g, often lessthan 700 m²/g and typically from about 500 to about 650 m²/g.

Another method of increasing the stability and/or acidity of the Yzeolites is by exchanging the Y zeolite with polyvalent metal cations,such as rare earth-containing cations, magnesium cations or calciumcations, or a combination of ammonium ions and polyvalent metal cations,thereby lowering the sodium content until it is as low as the valuesdescribed above after the first or second ammonium exchange steps.Methods of carrying out the ion exchange are well known in the art.

The catalyst used in the process disclosed herein is intended primarilyfor use as a replacement catalyst in existing commercial hydrocrackingunits. Its size and shape is, therefore, preferably similar to those ofconventional commercial catalysts. It is preferably manufactured in theform of a cylindrical extrudate having a diameter of from about 0.8-3.2mm ( 1/32-⅛ in). The catalyst can however be made in any other desiredform such as a sphere or pellet. The extrudate may be in forms otherthan a cylinder such as the form of a well-known trilobal or other shapewhich has advantages in terms or reduced diffusional distance orpressure drop.

Commercial hydrocracking catalysts contain a number of non-zeoliticmaterials. This is for several reasons such as particle strength, cost,porosity, and performance. The other catalyst components, therefore,make positive contributions to the overall catalyst even if not asactive cracking components. These other components are referred toherein as the support. Some traditional components of the support suchas silica-alumina normally make some contribution to the crackingcapability of the catalyst. In embodiments of the process andcomposition disclosed herein, the catalyst contains a relatively smallcontent of beta zeolite. The catalyst contains from 0.5 to 5 wt-%,preferably from 0.7 to 2.6 wt-%, of beta zeolite based on the combinedweight of the beta zeolite, Y Zeolite I (if any), Y Zeolite II, and thesupport all on a dried basis. As used herein, the weight on a driedbasis is considered to be the weight after heating in dry air at 500° C.(932° F.) for 6 hours. The catalyst has a weight ratio of the Y ZeoliteI to the beta zeolite of from 0.5 to 5, preferably from 0.5 to 2.0, on adried basis. When the optional Y Zeolite I is present, the catalyst hasa weight ratio of the Y Zeolite I to the Y Zeolite II of from 1.5 to 8,preferably from 2 to 6.5, on a dried basis. When the optional Y ZeoliteI is present, the catalyst contains from more than 5 wt-% to at most 15wt-%, of the Y Zeolite I and the Y Zeolite II based on the combinedweight of the beta zeolite, the Y Zeolite I, the Y Zeolite II, and thesupport, all on a dried basis.

The remainder of the catalyst particle besides the zeolitic material maybe taken up primarily by conventional hydrocracking materials such asalumina and/or silica-alumina. The presence of silica-alumina helpsachieve the desired performance characteristics of the catalyst. In oneembodiment the catalyst contains at least about 25 wt-% alumina and atleast about 25 wt-% silica-alumina, both based on the combined weight ofthe zeolites and the support, all on a dried basis. In anotherembodiment, the silica-alumina content of the catalyst is above about 40wt-% and the alumina content of the catalyst is above about 20 wt-%,both based on the combined weight of the zeolites and the support, allon a dried basis. However, the alumina is believed to function only as abinder and to not be an active cracking component. The catalyst supportmay contain over about 50 wt-% silica-alumina or over about 50 wt-%alumina based on the weight of the support on a dried basis.Approximately equal amounts of silica-alumina and alumina are used in anembodiment. Other inorganic refractory materials which may be used as asupport in addition to silica-alumina and alumina include for examplesilica, zirconia, titania, boria, and zirconia-alumina. Theseaforementioned support materials may be used alone or in anycombination.

Besides the beta zeolite, the Y zeolite, and other support materials,the subject catalyst contains a metallic hydrogenation component. Thehydrogenation component is preferably provided as one or more basemetals uniformly distributed in the catalyst particle. The hydrogenationcomponent is one or more element components from Groups 6, 9, and 10 ofthe periodic table. Noble metals such as platinum and palladium could beapplied but best results have been obtained with a combination of twobase metals. Specifically, either nickel or cobalt is paired withtungsten or molybdenum, respectively. The preferred composition of themetal hydrogenation component is both nickel and molybdenum or bothnickel and tungsten. The amount of nickel or cobalt is preferablybetween about 2 and about 8 wt-% of the finished catalyst. The amount oftungsten or molybdenum is preferably between about 8 and about 22 wt-%of the finished catalyst. The total amount of a base metal hydrogenationcomponent is from about 10 to about 30 wt-% of the finished catalyst.

The catalyst of the subject process can be formulated using industrystandard techniques. This can, with great generalization, be summarizedas admixing the beta zeolite and the Y zeolite with the other inorganicoxide components and a liquid such as water or a mild acid to form anextrudable dough followed by extrusion through a multihole die plate.The extrudate is collected and preferably calcined at high temperatureto harden the extrudate. The extruded particles are then screened forsize and the hydrogenation components are added as by dip impregnationor the well known incipient wetness technique. If the catalyst containstwo metals in the hydrogenation component these may be addedsequentially or simultaneously. The catalyst particles may be calcinedbetween metal addition steps and again after the metals are added.

In another embodiment, it may be convenient or preferred to combine theporous inorganic refractory oxide, the beta zeolite the Y zeolite, andcompound(s) containing the metal(s), then to comull the combinedmaterials, subsequently to extrude the comulled material, and finally tocalcine the extruded material. The comulling is effected with a sourceof metal, such as ammonium heptamolybdate or ammonium metatungstate andanother source of another metal, such as nickel nitrate or cobaltnitrate, with both source compounds generally being introduced into thecombined materials in the form of an aqueous solution or as a salt.Other metals can be similarly introduced in dissolved aqueous form or asa salt. Likewise, non-metallic elements, e.g., phosphorus, may beintroduced by incorporating a soluble component such as phosphoric acidsinto the aqueous solution when used.

Yet other methods of preparation are described U.S. Pat. No. 5,279,726and U.S. Pat. No. 5,350,501, which are hereby incorporated herein byreference in their entireties.

Catalysts prepared by the above-discussed procedures contain thehydrogenation metals in the oxide form. The oxide form is generallyconverted to the sulfide form for hydrocracking. This can beaccomplished by any of the well known techniques for sulfiding,including ex situ presulfiding prior to loading the catalyst in thehydrocracking reactor, presulfiding after loading the catalyst in thehydrocracking reactor and prior to use at an elevated temperature, andin situ sulfiding, i.e., by using the catalyst in the oxide form tohydrocrack a hydrocarbon feedstock containing sulfur compounds underhydrocracking conditions, including elevated temperature and pressureand the presence of hydrogen.

The hydrocracking process disclosed herein will be operated within thegeneral range of conditions now employed commercially in hydrocrackingprocesses. The operating conditions in many instances are refinery orprocessing unit specific. That is, they are dictated in large part bythe construction and limitations of the existing hydrocracking unit,which normally cannot be changed without significant expense, thecomposition of the feed and the desired products. The inlet temperatureof the catalyst bed should be from about 232° C. (450° F.) to about 454°C. (850° F.), and the inlet pressure should be from about 5171 kPa(g)(750 psi(g)) to about 24132 kPa(g) (3500 psi(g)), and typically fromabout 6895 kPa(g) (1000 psi(g)) to about 24132 kPa(g) (3500 psi(g)). Thefeed stream is admixed with sufficient hydrogen to provide a volumetrichydrogen circulation rate per unit volume of feed of about 168 to 1684normal ltr/ltr measured at 0° C. (32° F.) and 101.3 kPa(a) (14.7 psi(a))(1000 to 10000 standard ft³/barrel (SCFB) measured at 15.6° C. (60° F.)and 101.3 kPa(a) (14.7 psi(a))) and passed into one or more reactorscontaining fixed beds of the catalyst. The hydrogen will be primarilyderived from a recycle gas stream which may pass through purificationfacilities for the removal of acid gases although this is not necessary.The hydrogen rich gas admixed with the feed and in one embodiment anyrecycle hydrocarbons will usually contain at least 75 mol percenthydrogen. For hydrocracking to produce distillate the feed rate in termsof LHSV will normally be within the broad range of about 0.3 to 3.0hr-1. As used herein, LHSV means liquid hourly space velocity, which isdefined as the volumetric flow rate of liquid per hour divided by thecatalyst volume, where the liquid volume and the catalyst volume are inthe same volumetric units.

The typical feed to the process disclosed herein is a mixture of manydifferent hydrocarbons and coboiling compounds recovered by fractionaldistillation from a crude petroleum. It will normally have componentsthat boil higher than the upper end of the range of the 149° C. (300°F.) to 371° C. (700° F.) boiling range in order to produce distillate.Often it will have a boiling point range starting above about 340° C.(644° F.) and ending in one embodiment below about 482° C. (900° F.), inanother embodiment below about 540° C. (1004° F.), and in a thirdembodiment below about 565° C. (1049° F.). Such a petroleum derived feedmay be a blend of streams produced in a refinery such as atmospheric gasoil, coker gas oil, straight run gas oil, deasphalted gas oil, vacuumgas oil, and FCC cycle oil. A typical gas oil comprises components thatboil in the range of from about 166° C. (330° F.) to about 566° C.(1050° F.). Alternatively, the feed to the process disclosed herein canbe a single fraction such as a heavy vacuum gas oil. A typical heavy gasoil fraction has a substantial proportion of the hydrocarbon components,usually at least about 80 percent by weight, boiling from about 371° C.(700° F.) to about 566° C. (1050° F.). Synthetic hydrocarbon mixturessuch as recovered from shale oil or coal can also be processed in thesubject process. The feed may be subjected to hydrotreating or treatedas by solvent extraction prior to being passed into the subject processto remove gross amounts of sulfur, nitrogen or other contaminants suchas asphaltenes.

The subject process is expected to convert a large portion of the feedto more volatile hydrocarbons such as distillate boiling rangehydrocarbons. Typical conversion rates vary from about 50 to about 100volume-percent (hereinafter vol-%) depending greatly on the feedcomposition. The conversion rate is between from about 60 to about 90vol-% in an embodiment of the process disclosed herein, from about 70 toabout 90 vol-% in another embodiment, from about 80 and to about 90vol-% in yet another embodiment, and from about 65 to about 75 vol-% instill another embodiment. The effluent of the process will actuallycontain a broad variety of hydrocarbons ranging from methane toessentially unchanged feed hydrocarbons boiling above the boiling rangeof any desired product. The effluent of the process typically passesfrom a reactor containing a catalyst and is usually separated by methodsknown to a person of ordinary skill in the art, including phaseseparation or distillation, to produce a product having any desiredfinal boiling point. The hydrocarbons boiling above the final boilingpoint of any desired product are referred to as unconverted productseven if their boiling point has been reduced to some extent in theprocess. Most unconverted hydrocarbons are recycled to the reaction zonewith a small percentage, e.g. 5 wt-% being removed as a drag stream. Forproducing distillate, at least 30 wt-%, and preferably at least 50 wt-%,of the effluent boils below 371° C. (700° F.).

The process and composition disclosed herein can be employed in what arereferred to in the art as single stage and two stage process flows, withor without prior hydrotreating. These terms are used as defined andillustrated in the book titled Hydrocracking Science and Technology, byJ. Scherzer and A. J. Gruia, ISBN 0-8247-9760-4, Marcel Dekker Inc., NewYork, 1996. In a two-stage process the subject catalyst can be employedin either the first or second stage. The catalyst may be preceded by ahydrotreating catalyst in a separate reactor or may be loaded into thesame reactor as a hydrotreating catalyst or a different hydrocrackingcatalyst. An upstream hydrotreating catalyst can be employed as feedpretreatment step or to hydrotreat recycled unconverted materials. Thehydrotreating catalyst can be employed for the specific purpose ofhydrotreating polynuclear aromatic (PNA) compounds to promote theirconversion in subsequent hydrocracking catalyst bed(s). The subjectcatalyst can also be employed in combination with a second, differentcatalyst, such as a catalyst based upon Y zeolite or having primarilyamorphous cracking components.

In some embodiments of the process disclosed herein, the catalyst isemployed with a feed or in a configuration that the feed passing throughthe catalyst is a raw feed or resembles a raw feed. The sulfur contentof crude oil, and hence the feed to this process, varies greatlydepending on its source. As used herein, a raw feed is intended to referto a feed which has not been hydrotreated or which still containsorganic sulfur compounds which result in a sulfur level above 1000wt-ppm or which still contains organic nitrogen compounds that result ina nitrogen level above 100 wt-ppm (0.01 wt-%).

In other embodiments of the process disclosed herein, the catalyst isused with a feed that has been hydrotreated. Persons of ordinary skillin the art of hydrocarbon processing know and can practice hydrotreatingof a raw feed to produce a hydrotreated feed to be charged to theprocess disclosed herein. Although the sulfur level of the hydrotreatedfeed may be between 500 and 1000 wt-ppm, the sulfur level of thehydrotreated feed is less than 500 wt-ppm in one embodiment of theprocess disclosed herein and from 5 to 500 wt-ppm in another embodiment.The nitrogen level of the hydrotreated feed is less than 100 wt-ppm inone embodiment and from 1 to 100 wt-ppm in another embodiment.

All references herein to the groups of elements of the periodic tableare to the IUPAC “New Notation” on the Periodic Table of the Elements inthe inside front cover of the book titled CRC Handbook of Chemistry andPhysics, ISBN 0-8493-0480-6, CRC Press, Boca Raton, Fla., U.S.A.,80^(th) Edition, 1999-2000. All references herein to surface area are tosingle-point surface areas at a nitrogen partial pressure of p/p_(o) of0.03 as determined by the BET (Brunauer-Emmett-Teller) method usingnitrogen adsorption technique as described in ASTM D4365-95, StandardTest Method for Determining Micropore Volume and Zeolite Area of aCatalyst, and in the article by S. Brunauer et al., J. Am. Chem. Soc.,60(2), 309-319 (1938). All references herein to boiling points are toboiling points as determined by ASTM D2887, Standard Test Method forBoiling Range Distribution of Petroleum Fractions by Gas Chromatography.ASTM methods are available from ASTM International, 100 Barr HarborDrive, P.O. Box C700, West Conshohocken, Pa., U.S.A.

The following examples are provided for illustrative purposes and not tolimit the process and composition as defined in the claims.

EXAMPLE 1 Sample 1

A modified Y zeolite was prepared by steaming an ammonium exchanged Yzeolite sold by UOP LLC (Des Plaines, Ill., USA) and referred to in theliterature as Y-84 having a sodium content of less than about 0.2 wt-%calculated as Na₂O. The resulting modified Y zeolite is referred toherein as Sample 1 and had an overall silica to alumina (SiO₂ to Al₂O₃)mole ratio of 5.0 to 5.5, a unit cell size of 24.28 angstrom, and asurface area of 540 to 640 m²/g. Sample 1, which is an example of YZeolite I, is referred to in the Table as Y1.

Sample AW 1

A sample of Sample 1 was acid washed. The resulting acid-washed modifiedY zeolite is referred to herein as Sample AW 1 and had an overall silicato alumina (SiO₂ to Al₂O₃) mole ratio of 11.0, a unit cell size of 24.28angstrom, and a surface area of 570 to 750 m²/g. Sample AW 1, which isan example of Y Zeolite I, is referred to in the Table as AW Y1.

Sample 2

A modified Y zeolite was prepared in a manner similar to that describedfor Sample 1, except the steaming conditions were different. Theresulting modified Y zeolite is referred to herein as Sample 2 and hadan overall silica to alumina (SiO₂ to Al₂O₃) mole ratio of from 5.0 to5.5, a unit cell size of 24.35 angstrom, and a surface area of 630 to730 m²/g. Sample 2, which is an example of Y Zeolite II, is referred toin the Table as Y2.

EXAMPLE 2

Eight catalysts (A-H) were prepared by mixing Sample 1 if present,Sample AW 1 if present, Sample 2 if present, a beta zeolite having anoverall silica to alumina (SiO₂ to Al₂O₃) mole ratio of 23.8 and an SF₆adsorption capacity of 29 wt-% if present, amorphous silica-alumina, andHNO₃-peptized Catapal™ C boehmite alumina in a muller. The beta zeolite,which had an overall silica to alumina (SiO₂ to Al₂O₃) mole ratio of23.8 and an SF₆ adsorption capacity of 29 wt-%, either contained thetemplate used during its synthesis or had been subsequently calcined atmild conditions to remove the aforementioned template. In the Table, thebeta zeolite containing the template is referred to as Beta 1, and thecalcined beta zeolite is referred to as Beta 2. The amorphoussilica-alumina was either CCIC silica-alumina which had a nominalcomposition of 75 wt-% silica and 25 wt-% alumina, or Siral 40silica-alumina, which had a nominal composition of 40 wt-% silica and 60wt-% alumina. CCIC silica-alumina is available from Catalysts &Chemicals Industries Co. Ltd. (CCIC), and Catapal C alumina and Siral 40silica-alumina are available from Sasol Germany GmbH. The amounts ofthese components on a dried basis in each final catalyst are listed inthe Table. The resulting mixture was extruded into 1.6 mm ( 1/16 in)diameter cylindrical particles of between 3.2 mm (⅛ in) and 12.7 mm (½in) in length. The wet extrudates were dried at 104° C. (220° F.) for aminimum of 4 hr and then calcined at temperatures in excess of 550° C.(1022° F.) for a minimum of 90 minutes. For catalysts A-F and H,sufficient nickel nitrate to provide 4 wt-% nickel (calculated as Ni) inthe final catalyst and sufficient ammonium metatungstate to provide 14wt-% tungsten (calculated as W) in the final catalyst were then added tothe calcined extrudates to incipient wetness, while for catalysts G andH the corresponding amounts were 5 wt-% nickel and 17.5 wt-% tungsten.The extrudates were then dried to be free-flowing, and then oxidized bycalcining at about 500° C. (932° F.) for a minimum of 90 minutes.Catalyst I is a standard hydrocracking catalyst containing on average5.5 wt-% nickel and 17.5 wt-% tungsten. It is believed that thedifferences in nickel and tungsten contents do not have a significanteffect on the hydrocracking activity and selectivity results describedin these examples.

EXAMPLE 3

Each of the above-described nine catalysts was pre-sulfided by passing agas stream consisting of 10 vol-% H₂S and the balance H₂ through a bedof the catalyst at a temperature initially of about 149° C. (300° F.)and slowly raised to 413° C. (775° F.) and held at the temperature forabout 6 hours.

The nine catalysts were compared for hydrocracking activity andselectivity (i.e., product yields) in simulated first stage testing.Specifically, the nine catalysts were separately tested forhydrocracking a hydrotreated light. Arabian vacuum gas oil (VGO) feedhaving a specific gravity of 0.877 at 15.6° C. (60° F.) (API gravity of30.05°), an initial boiling point of 107° C. (224° F.), a 5 wt-% boilingpoint of 195° C. (382° F.), a final boiling point of 550° C. (1021° F.),and a 50 wt-% boiling point of 24° C. (795° F.), with about 13 wt-%boiling below 288° C. (550° F.) and about 26 wt-% boiling below 371° C.(700° F.).

Each catalyst was tested for simulated first stage operation by passingthe feedstock through a laboratory size reactor at a LHSV of 1.5 hr-1, atotal pressure of 13786 kPa(g) (2000 psi(g)), and a volumetric hydrogenfeed rate per unit volume of feed of 1684 normal ltr/ltr measured at 0°C. (32° F.) and 101.3 kPa(a) (14.7 psi(a)) (10000 SCFB measured at 15.6°C. (60° F.) and 101.3 kPa(a) (14.7 psi(a))). Sufficient di-tert-butyldisulfide was added to the feed to provide 2.1 wt-% sulfur and therebyto simulate a hydrogen sulfide-containing atmosphere as it exists incommercial first stage hydrocracking reactors. In addition, sufficientcyclohexylamine was added to the feed to provide 780 wt-ppm nitrogen andthereby to simulate an ammonia-containing atmosphere as it exists incommercial first stage hydrocracking reactors.

For hydrocracking tests to produce distillate, the temperatureconditions were adjusted as necessary to maintain about a 65 wt-% netconversion to materials boiling below 371° C. (700° F.), over the courseof 100 hours. Net conversion is the effluent boiling below 371° C. (700°F.) as a percentage of the feed minus the percentage of the feed boilingbelow 371° C. (700° F.). At the end of the 100 hours, the temperaturerequired to maintain the 65 wt-% net conversion was recorded, and therelative activities and selectivities of each catalyst were calculated.These data are summarized in the Table. The selectivity values for eachcatalyst were total distillate (i.e., 149° C. (300° F.) to 371° C. (700°F.)), light distillate (i.e., 149° C. (300° F.) to 288° C. (550° F.)),and heavy distillate (i.e., 288° C. (550° F.) to 371° C. (700° F.)). Therelative activity value for each catalyst is entered as the differencebetween the required temperature of the catalyst to maintain the 65 wt-%net conversion and a reference temperature that was the same for allnine catalysts. The lower the value for relative activity, the moreactive is the catalyst.

TABLE Catalyst Designation A B C D E F G H I Composition, wt-% ofzeolite and support Zeolites 5.4 6.1 7.2 10.0   6.75 8.0 20.0 21.0 10.0Y zeolite 3.2 3.5 6.0 8.5 6.0 6.5 20.0 21.0 10.0 Y1 Zeolite — — 4.5 7.0— — 20.0 12.0 10.0 AW Y1 Zeolite — — — — 5.0 5.0 — — — Y2 Zeolite 3.23.5 1.5 1.5 1.0 1.5 —  9.0 — Beta zeolite 2.2 2.6 1.2 1.5  0.75 1.5 — —— Beta 1 Zeolite 2.2 2.6 1.2 1.5 — — — — — Beta 2 Zeolite — — — —  0.751.5 — — — Support 94.6  93.9  92.8  90.0  93.25 92.0  80.0 79.0 90.0CCIC silica-alumina — 60.0  60.0  60.0  60.0  60   50.0 50.0 43.5 Siral40 silica-alumina 70.0  — — — — — — — — Alumina 24.6  33.9  32.8  30.0 33.25 32.0  30.0 29.0 46.5 Zeolite ratios, wt/wt Y2:Beta 1.5 1.3 1.3 1.01.3 1.0 NA NA NA (Y1 + AW Y1 + Y2):Beta 1.5 1.3 5.0 5.7 8.0 4.3 NA NA NA(Y1 + AW Y1):Y2 NA NA 3.0 4.7 5.0 3.3 NA 1.3 NA 149° C. (300° F.) to371° C. (700° F.) cut Relative activity, ° C. (° F.) +5.0  +4.2  +8.9 +5.3  +11.4  +6.7  +8.3 +1.1 +13.9  (+9.0)  (+7.5)  (+16)    (+9.5) (+20.5)  (+12)    (+15)   (+2.0) (+25)   Yield, wt-% 51.5  50.9  52.2 51.5  52.4  51.8  49.9 49.1 51.2 149° C. (300° F.) to 288° C. 34.2  34  34.3  34.4  34.7  34.5  34.1 34.4 34.5 (550° F.) cut yield, wt-% 288° C.(550° F.) to 371° C. 17.3  16.9  17.9  17.1  17.7  17.3  15.8 14.7 16.7(700° F.) cut yield, wt-% Yield ratio: (288° C. to 371° C.  0.51  0.50 0.52  0.50  0.51  0.50  0.46  0.43  0.48 cut):(149° C. to 288° C. cut)NA = Not Applicable

FIG. 1 is a chart of the 149° C. (300° F.) to 371° C. (700° F.) cutdistillate selectivity of Catalysts A-I plotted versus the relativecatalyst activity expressed in terms of reactor temperature above thereference temperature required to achieve 65 wt-% net conversion of theVGO to the total distillate cut. Catalysts A-F (squares) show more totaldistillate selectivity at a given relative activity than Catalysts G-I(diamonds). FIG. 2 is a chart of the weight ratio of the heavydistillate cut selectivity to the light distillate cut selectivityversus the relative activity. Catalysts A-F (squares) show asignificantly higher selectivity of heavy distillate relative to lightdistillate compared to Catalysts G-I (diamonds).

1. A process for hydrocracking a hydrocarbon feed stock which comprisescontacting the feedstock at a temperature from about 232° C. to about454° C. and at a pressure from about 5171 kPa(g) to about 24132 kPa(g)in the presence of hydrogen with a catalyst comprising a hydrogenationcomponent, a beta zeolite having an overall silica to alumina mole ratioof less than 30 and a SF₆ adsorption capacity of at least 28 wt-%, a Yzeolite having a unit cell size of from 24.33 to 24.38 angstrom (YZeolite II), and a support, wherein the Y Zeolite II has an overallsilica to alumina mole ratio of from 5.0 to 1.0, wherein the catalystcontains from 0.5 to 5 wt-% beta zeolite based on the combined weight ofthe beta zeolite, the Y Zeolite II, and the support on a dried basis,and wherein the catalyst has a weight ratio of the Y Zeolite II to thebeta zeolite of from 0.5 to 5 on a dried basis.
 2. The process of claim1 wherein the Y Zeolite II has a surface area of less than 800 m²/g. 3.The process of claim 1 wherein the Y Zeolite II is prepared by a processcomprising the steps of: a) partially ammonium exchanging a sodium Yzeolite; b) calcining the zeolite resulting from step (a) in thepresence of water vapor; c)ammonium exchanging the zeolite resultingfrom step (b); and d) calcining the zeolite resulting from step (c) inthe presence of water vapor.
 4. The process of claim 1 wherein the YZeolite II is prepared by a process comprising the steps of: a)partially ammonium exchanging a sodium Y zeolite; b) calcining thezeolite resulting from step (a) in the presence of water vapor; c)contacting the zeolite resulting from step (b) with a fluorosilicatesalt in the form of an aqueous solution; and d) calcining the zeoliteresulting from step (c) in the presence of water vapor.
 5. The processof claim 1 wherein the Y Zeolite II is prepared by a process comprisingthe steps of: a) contacting a sodium Y zeolite with a fluorosilicatesalt in the form of an aqueous solution; and b) calcining the zeoliteresulting from step (a) in the presence of water vapor.
 6. The processof claim 1 wherein the hydrogenation component is selected from thegroup consisting of molybdenum, tungsten, nickel, cobalt, and the oxidesand sulfides thereof.
 7. The process of claim 1 wherein the Y Zeolite IIhas a unit cell size of from 24.34 to 24.36 angstrom.
 8. The process ofclaim 1 wherein the weight ratio of the Y Zeolite II to the beta zeoliteis from 0.5 to 2.0 on a dried basis.
 9. The process of claim 1 whereinthe unit cell size of the Y Zeolite II is a first unit cell size, andthe catalyst comprises an additional Y zeolite (Y Zeolite I) having asecond unit cell size at least 0.04 angstrom smaller than the first unitcell size.
 10. The process of claim 9 wherein the second unit cell sizeis from 24.25 to 24.32 angstrom.
 11. The process of claim 9 wherein thecatalyst has a weight ratio of the Y Zeolite I to the Y Zeolite II offrom 1.5 to 8 on a dried basis.
 12. The process of claim 9 wherein thecatalyst contains from more than 5 wt-% to at most 15 wt-% of the YZeolite I and the Y Zeolite II based on the combined weight of the betazeolite, the Y Zeolite I, the Y Zeolite II, and the support on a driedbasis.
 13. The process of claim 9 wherein the Y Zeolite I is prepared bya process comprising the steps of: a) partially ammonium exchanging asodium Y zeolite; b) calcining the zeolite resulting from step (a) inthe presence of water vapor; c)ammonium exchanging the zeolite resultingfrom step (b); and d) calcining the zeolite resulting from step (c) inthe presence of water vapor.
 14. The process of claim 9 wherein the YZeolite I is prepared by a process comprising the steps of: a) partiallyammonium exchanging a sodium Y zeolite; b) calcining the zeoliteresulting from step (a) in the presence of water vapor; c) contactingthe zeolite resulting from step (b) with a fluorosilicate salt in theform of an aqueous solution; and d) calcining the zeolite resulting fromstep (c) in the presence of water vapor.
 15. The process of claim 9wherein the Y Zeolite I is prepared by a process comprising the stepsof: a) contacting a sodium Y zeolite with a fluorosilicate salt in theform of an aqueous solution; and b) calcining the zeolite resulting fromstep (a) in the presence of water vapor.
 16. A composition of mattercomprising a catalyst comprising a hydrogenation component, a betazeolite having an overall silica to alumina mole ratio of less than 30and a SF₆ adsorption capacity of at least 28 wt-%, a Y zeolite having aunit cell size from 24.33 to 24.38 angstrom (Y Zeolite II), and asupport, wherein the Y Zeolite II has an overall silica to alumina moleratio of from 5.0 to 11.0, wherein the catalyst contains from 0.5 to 5wt-% beta zeolite based on the combined weight of the beta zeolite, theY Zeolite II, and the support on a dried basis, and wherein the catalysthas a weight ratio of the Y Zeolite II to the beta zeolite of from 0.5to 5 on a dried basis, wherein the Y Zeolite II has a surface area ofless than 800 m²/g.
 17. The composition of claim 16 wherein thehydrogenation component is selected from the group consisting ofmolybdenum, tungsten, nickel, cobalt, and the oxides and sulfidesthereof.
 18. The composition of claim 16 wherein the unit cell size ofthe Y Zeolite II is a first unit cell size, and the catalyst comprisesan additional Y zeolite (Y Zeolite I) having a second unit cell size offrom 24.25 to 24.32 angstrom and at least 0.04 angstrom smaller than thefirst unit cell size.
 19. A hydrocracking process comprising contactinga hydrocarbon feedstock with a catalyst at a temperature between about232° C. and about 454° C. and at a pressure between about 5171 kPa(g)and about 24132 kPa(g) in the presence of hydrogen so as to produce aneffluent of lower average boiling point than the hydrocarbon feedstock,the catalyst comprising one or more hydrogenation components incombination with a support comprising an inorganic refractory oxide,zeolite beta in a form catalytically active for cracking hydrocarbons, aY zeolite catalytically active for cracking hydrocarbons and having afirst unit cell size of from 24.33 to 24.38 angstrom (Y Zeolite II),wherein the Y Zeolite II has an overall silica to alumina mole ratio offrom 5.0 to 11.0 and a surface area of less than 800 m²/g, and anadditional Y zeolite (Y Zeolite I) having a second unit cell size offrom 24.25 to 24.32 angstrom, wherein the second unit cell size is atleast 0.04 angstrom smaller than the first unit cell size, wherein thecatalyst contains from 0.5 to 5 wt-% beta zeolite based on the combinedweight of the beta zeolite, the Y Zeolite I, the Y Zeolite II, and thesupport on a dried basis, wherein the catalyst has a weight ratio of theY Zeolite II to the beta zeolite of from 0.5 to 5 on a dried basis, andwherein the catalyst has a weight ratio of the Y Zeolite I to the YZeolite II of from 1.5 to 8 on a dried basis.
 20. The hydrocrackingprocess of claim 19 wherein at least 30 wt-% of the effluent boils below371° C.